Heat transfer apparatus



March 1956 N. DICKINSON HEAT TRANSFER APPARATUS 2 Sheets-Sheet 1Original Filed Aug. 2'7, 1949 -.m 5:5: mm\ m mm m0. .u um mm zuwomocizoifiiifiwwaomo: 5 3 h 102298;. 555m m mm m 3 ozimouom z f wv AV A hm VE. A, a T v s3 v n R mm a w... A A Emmomm/L N\ 626%. h L mm wmmm g Al mm3 mm 0| a lo N 5 5 J on 2. Q.

llllitll I INVENTOR. NORMAN L. DICIHNJON ATTORNEYS March 27, 1956 N. IDICKINSON 2,739,880

HEAT TRANSFER APPARATUS $52 f HYDROGEN INVENTOR. N RMAN L. DICKINSONATTO ZNEYS between the inlet and outlet of the reactor.

HEAT TRANSFER APPARATUS Norman L. Dickinson, Basking Ridge, N. 1., assignor to The M. W. Kellogg Company, Jersey City, N. 1., a corporationof Delaware Original application August .27, 1949, Serial No. 112,777,

now Patent No. 2,642,381, dated June 16, 1953. Divided and thisapplication December 21, 1951, Serial No. 262,849

5 Claims. (Cl. 23-288) The invention relates to heat transfer between anexothermic reaction and an endothermic reaction. In one aspect theinvention relates to a method for transferring heat from a relativelylow temperature exothermic reaction to a relatively high temperatureendothermic reaction. In another aspect the invention relates to anintefraction. In this example, the 'hydroforming reaction isendothermic, which reaction requires the addition of heat in order tomaintain substantially isothermal temperature conditions. Although theprior hydrodesulfurization reaction is exothermic with the liberation ofa considerable amount of heat, the hydrodesulfurization reaction iseffected at the same or lower temperature than the subsequenthydroforming reaction. This characteristic of the two reactionsprohibits direct heat exchange between the reaction zones so as toutilize the exothermic heat of reaction of the hydrodesulfurizationstep. The heat to the second or hydroforming step may be supplied byindirect :heat exchange with a relatively high temperature fluidexternally heated, or by the more usual manner ofsuperheating the feedto the hydroforming step. In :general, it may be said that either of theabove methods are undesirable in an integrated process since they do notutilize the exothermic heat of reaction liberated in the.hydrodesulfuriz'ation step. Superheating has the disadvantageob-causing thermal decomposition of the feed. Supplying the endothermicheat of reaction by superheat is characterized by an-urrdesirably largetemperature gradient Thus, .the inlet temperature is above and theoutlet is below the optimum for the reaction. It is much to be desired,therefore, to provide a method for utilizing the exothermic heat ofreaction of the hydrodesulfurization stepin the endothermic hydroformingstep and also to provide -a method for at least minimizing thetemperature to which the feed to the hydroforming step must bepreheated.

It is an object of this invention to provide a method and apparatus fortransferring heat from an exothermic reaction to an endothermic reactionin an integrated process in which the exothermic reaction is carried"out at a temperature not higher than the temperature of the endothermicreaction.

Another object of this invention is to provide .an improved method fortreating a hydrocarbon-distillate stock.

Another object of this invention isto provide an proved method for thehydrodesulfurization"Land.fhydroforming'of a distillate.

It is a further object of this invention to ;.provide a Sttes Patent2,73%,880 Patented Mar. 27, 1956 2 process for the removal of nitrogenand sulfur compounds from a hydrocarbon distillate.

Another object of this invention is to provide a novel arrangement ofapparatus for hydrodesulfurization and hydroforming.

Still another object of this invention is to provide a novel fluidizedtype of operation for the hydrodesulfurization of a hydrocarbondistillate.

Another object of this invention is to provide a novel type of fluidizedprocess for the hydroforming of a hydrocarbon distillate.

It is a further object of this invention to provide a process forincreasing the catalyst life in a hydrodesul- .f-urization process andin a hydroforming process.

Various other objects and advantages will become apparent to thoseskilled in the art from the accompanying description and disclosure.

This invention applies specifically to an integrated process involvingboth exothermic and endothermic reaction steps on a common feed stock,said steps being effected at substantially the same temperature, or oneof said steps being a relatively low temperature exothermic reaction andanother of said steps being a relatively high temperature endothermicreaction. According to this invention,

' exothermic heat is made available for the endothermic reaction byvaporizing a liquid in indirect heat exchange with the exothermicreaction, compressing the vapors thus produced, thereby increasing theirtemperature and raistionary or fixed bed in which catalyst is containedin tubes surrounded'by a heat exchange fluid, or in a fluidizedcondition, or even in the absence of a catalyst. The invention .hasparticular applicability to a system in which reactant gases are passedat a relatively high velocity through a reaction zone in the presence offinely divided catalyst under conditions such that the catalyst issuspended or entrained in the gaseous reaction mixture. The inventionmay be best understood by its application .to .a .specific integratedprocess and the inventor has chosentor the description of this inventionits application to an integrated process involving the successive stepsof hydrodesulfurization and hydroforming of a sulfurbearing.hydrocarbondistillate. It is to be understood, however, that the invention haswider application to exothermic and endothermic reactionsgenerallyjinwhich the exothermic heat is available at the same or lowertem- :and various arrangements of'apparatussuitable for carrying out thepresent invention. Figure l ofthe drawings is .an arrangementofapparatus and flow diagram of an integrated processior thehydrodesulfurization and-hydroforming of .a hydrocarbon distillatefraction containing v:organically combinedlsulfur. In Figure .1, asingle heat exchange means is provided in each reactionzone. ,Fig-

.ure .2. represents .a:diagrammatic illustration. of .a heat tex- -change..system similarto that "shown in Figure ,1 except that multipleheat-exchangers are employed in eachof thetreaction zones. Figure 3diagrammatically illustrates by a pump, not shown.

naphtha of the following characteristics are treated:

. Table 1 Gravity, API 46.0 A. S. T. M. distillation:

I. B. P., F 217 5% 241 10 249 20 261 30 272 40 284 50 298 60 313 70 33080 348 90 373 95 395 E. P 426 Recv., percen 98.0 Residue 1.3 Loss 0.7Aniline point, F 89 Octane number, CFRM 69.4 Reid vapor pressure, p. s.i 0.6 Sulfur, wt. percent 2.2 Nitrogen, wt. per 0.20

Although Figure l is described with respect to a specific feed stock,the invention may be applied to the hydrodesulfurization of any straightrun or cracked naphtha containing from about 0.05 to about weight percent sulfur and boiling within the range of about 100 to about 600 F.Preferably, the feed stock to be treated in accordance with a process ofthe type shown in Figure 1 boils within the range of about 250 F. toabout 500 F. and has and A. P. I. gravity between about 40 and about 60.in treating such a feed stock according to the process to be hereinafterdescribed, 90 to about 100 per cent of the organic sulfur is convertedto hydrogen sulfide and removed.

The aforesaid quantity of naphtha is introduced into the system throughconduit 4 and it is admixed with about 62,500,000 standard cubic feetper day of hydrogen-containing gas which includes 7,500,000 standardcubic feet of make-up hydrogen from conduit 6 and 55,000,000 standardcubic feet of hydrogen-containing recycle gas from conduit 37 or 57. Themixture comprising hydrogen and naphtha is then passed through conduit 4to a preheater and vaporizer 7 which preheats the mixture to atemperature substantially equivalent to the desired hydrodesulfurizationtemperature which in this example is 850 F. The feed stock is pumped tothe desired pressure Since the process of Figure 1 involves bothdesulfurization and hydroforming and utilizes catalyst transfer betweeneach of the aforesaid zones, it is preferred to effect both reactions atsubstantially the same pressure in order to minimize difficulties in thetransfer of catalyst between zones. In the process illustrated, thepressure is 500 pounds per square inch gage.

From preheater 7, the vaporized feed stock is passed through conduit 8to hydrodesulfurization reactor 9. Catalyst is introduced into transferline 8 through a conduit or standpipe 47. This catalyst is a finelydivided solid material, the nature and composition of which will bediscussed more fully hereinafter. The catalyst is suspended in thereaction gases and passes by entrainment to reactor 9-. In reactor 9,the naphtha vapors pass upward through an elongated and enlargedreaction zone at a sufiiciently high gas velocity that the catalystparticles are suspended and continuously move in the direction of flowof the gases in concurrent flow. The velocity of the gases should beabove about 5 feet per second and may be as high as 40 or 50 feet persecond in order to maintain the catalyst particles moving in thedirection of flow of the gases and to prevent the formation of theconventional pseudoliquid dense phase of catalyst particlescharacterized by the internal circulation of catalyst particles from topto bottom of the dense phase. Normally, the higher the velocity the moreuniform is the flow of catalyst and gases. Thus, the higher-velocitiesare preferred from the standpoint of the physical condition of thereaction mixture. However, higher velocities require longer reactionchambers, and, from a construction standpoint, lower velocities withinthe above range are preferable. The preferred velocities, taking intoaccount both the physical condition of the reaction mixture and thelength of the reaction chamber, are between about 6 and about 15 feetper second. Employing such relatively high velocities, the catalyst isentrained in a highly dispersed condition in the gaseous reactants inreactor 9. The concentration of the contact material under suchconditions of velocity and a suitable loading rate is between about 1and about 20 pounds per cubic foot. In comparison with the conventionalpseudo-liquid dense phase of finely divided contact material, thisconcentration is one-half or less of the density or concentration ofcatalyst in the dense phase.

At a space velocity of about 7 w./hr./w. (weight of naphtha per hour perweight of catalyst) and a reactor 3 feet 3 inches in inside diameter andfeet in length, a gas velocity of about 6.8 feet per second and acatalyst concentration in reactor 9 of about 25 pounds per cubic footare suitable for the operation to achieve the desired conversion. In theoperation shown in Figure 1, to per cent of the sulfur is converted tohydrogen sulfide, in the presence of a suitable catalyst. Nitrogenousorganic compounds are also reduced to ammonia and hydrocarbons. The morereactive unsaturated compounds,

such as mono-olefins and diolefins, are simultaneously hydrogenated totheir corresponding saturated compounds. In the particular reactionillustrated in Figure 1, about 16,700,000 B. t. u.s per hour of heat arereleased at a substantially isothermal temperature of about 850 F.

'An efiiuent comprising desulfurized naphtha, hydrogen sulfide,hydrogen, relatively low boiling hydrocarbons and entrained catalyst ispassed from reactor 9 through conduit 11 to separator 12. The effluentis introduced through conduit 11 tangentially into separator 12 to aidin the separation of solid contact material from the reaction efiluent.A cyclone separator (not shown) may be positioned within separator 12 toaid in removal of catalyst dust from the reaction efiluent. The cycloneseparator contains a dip leg projecting into the lower portion ofseparator 12. Contact material substantially at the temperature ofreactor 9 settles to the bottom of separator 12 and passes through afunnel shaped bottom portion 13 and through a standpipe or conduit 14into a stripping zone 15 which may contain bafiles or the like. Thelevel of the catalyst bed in the stripping zone 15 is above the lowerend of standpipe 14 as indicated by numeral 17. Stripping gas isintroduced into the lower portion of stripper 15 through a conventionalinlet conduit and distribution means 18. The stripping gas along withstripped material is withdrawn from the upper portion of stripper 15through conduit 16 or may be passed directly into separator 12 by meansnot shown. The stripping gas removes tarry and carbonaceous deposits andoccluded gases from the catalyst and may also condition the catalyst. Asuitable stripping gas comprises any of the following: recycle gas,hydrogen, methane, carbon dioxide, steam, air, and nitrogen. Thestripped catalyst is withdrawn from stripper 15 through conduit 19 andis passed to transfer line 39 of hydroforming reactor 41. In amodification of the invention, a portion or all of the catalyst may berecycled through conduit 20 to reactor 9. In general, it is preferableto recycle at least a portion of the stripped catalyst directly toreactor 9 in order to control theconcentration of catalyst therein.independent "bf c atalyst circulation throughout the unit. The ratio ofrecycle catalyst to catalyst passed to reactor 41 is usually about :1.

The effluent from reactor 9, as previously discussed, is removed fromthe upper part of separator 12 and is sub- 'stantially free fromentrained or suspended catalyst. However, in order to remove the lasttraces of suspended catalyst, this effluent is passed through metallicfilters (not shown) or other known means of dust recovery and thenpassed through conduit 21, condenser 22, conduit 23 to accumulator 24.Condenser 22 represents a single or series of condensation units at thesame or successively lower pressures and temperatures. Preferably,operating pressure is employed on the condensation unit 22 andaccumulator 24. The temperature of the effluent after passage throughcondensing unit 22 is about 100 F. In accumulator 24, the desulfurizednaphtha is collected and passed through conduit 26 to separating unit 27which represents conventional equipment for the fractionationandseparanon of a naphtha suitable for use in the subsequenthydroforming-step. This naphtha stream is free of C hydrocarbons andlighter, preferably free of Ce hydrocarbons and lighter. Ordinarily,separating unit 27 is of conventional design suitable for removal 'ofthe dissolvdhydrogen sulfide, ammonia and relatively low boilfnghydrocarbons, such as ethane and propane. These gases are removed fromseparating unit 27 through eon- 'dui't 28. Desulfurized naphtha iswithdrawn from separating unit 27 through conduit 29 and is introducedinto transfer line 31 of hydroforming reactor '41. The desulfurizednaphtha in this example has approximately the following properties:

I. B. P., F 143 'Reeov., per cent 98.0

Residue 1.0

Aniline point, F 120 Octane number, CFRM' 55 Reid vapor pressure, p. s.i 1.5

-Sulfu'r, wt. per cent 0.04

Nitrogen, wt. percent 0.07

"medium is withdrawn from extractor 33 through conduit -36 and passed toa conventional stripping or desorption unit (not shown) in which thehydrogen sulfide and aminonia are stripped from the absorption medium.The stripped absorption medium is thereafter returned to the upperportion of extraction unit 33 through inlet conduit 34. The gaseouseffiuent from extraction unit 33 comprises hydrogenand methane, ethaneand propane. This =1'gaseous stream is-r'ecycled through conduit -37 toconduit "4 in order to supply the excess-hydrogen'tothe'hydrodesulfurization reaction. In theexample described in Figure a1, the quantities of net materials, excluding recycle from Desulfurizednaphtha (conduit 29)--l0, 220 barrels per stream day If desired, aportion of the gas in conduit 37 may be bypassed to conduit 31 leadingto hydroforming unit 41, as shown. In most instances, however, all ofthe gas through conduit 37 is returned to conduit 4.

The desulfurized naphtha is now hydrofo'rmed to improve the octane valuethereof by increasing its aromatic content. This is accomplished'bypassing liquid desulfurized naphtha through conduit 29 to conduit 31 andthence to a heating and vaporizing zone 38. Recycle gas from thehydroforming unit 41 containing hydrogen and relatively low boilinghydrocarbons is introduced into this stream prior to vaporizer 38through conduits 53 and 54, as shown. In vaporizer 38, the naphtha isvaporized and preheated to about the hydroforming temperature employedin hydroforming reactor 41. If de sired, the recycle gases in conduit 53may be introduced into the naphtha feed line after vaporizer 38 by meansnot shown. In such a case an additional preheater, not shown, must beincluded on recycle line 53.

The quantity of recycle gas is approximately 51,000,000 standard cubicfeet per day. The mixture of vaporized naphtha and recycle gaspasses ata relatively high velocity through conduit or transfer line 39 tohydroforming reactor 41 of enlarged cross section similar in design toreactor 9. Catalyst is picked up in conduit 39 from conduit or standpipe19. In reactor 41 the gas velocity is between about 5 and about 50 feetper second in order to entrain the catalyst particles in a relativelydispersed condition in the gaseous reactants. The physical condition ofthe catalyst andgases in hydroforming reactor' ll is substantially thesame as that described with respect to hydrodesulfurization reactor 9.

At the temperature of reaction of about 925 F. and a pressure of about480 pounds per square inch gage for hydroforming reactor 41, theprincipal reaction 'isthe dehydrogenation of n'aphthenic hydrocarbons toaromatics. To a somewhat lesser extent aliphatic hydrocarbons aredehydrogenated and cyclicized to aromatic compounds. Some cracking alsooccurs, but the hydrogen pressure is s'ufiiciently high to saturatesubstantially all of the cracked products and to substantially suppresspolymerization reactionsleading to compounds hig her boiling than thefeed. The combined elfect of these reactions is endothermic. Theexternal heat r'equired'to maintain substantially isothermal conditionsof approximately 925 F. is 16,200,000 B. t. u.s per hour.

The gaseous effluent containing entrained finely divided contactmaterial is withdrawn from hydroforming reactor 41 through conduit 42and introduced tangentially into separator 43. In separator 43, thecontact material is separated from the gaseous efiluen't and'is'p'assedinto a stripping zone 44. In stripping zone 44, tarry and carbonaceousdeposits and occluded gases are removed from the contact material.Stripping gas is'introt'luced into stripper 44 through conduit-46 andcomprises similar gases as described with respect to stripper 1'5 'her'einbefore described. The construction and operation of separator 43and stripper 44 aresubs'tan'tially the same as described with respect'toseparator 12 and stripper 15. Stripped catalyst is removed from stripper44 through conduit 47 and is passed to transfer line 8 from where it isreturned to hydrodesulfurization reactor 9 and the 'cycle'repeated.However, a'p'ortio'n or all o'f'the catalyst may be passedthroughconduit or-standpipe48 forrecycle tohydroforming unit 41 in'order to control the concentration of catalyst therein independent "ofcatalyst assassin circulation throughout the entire two stages of theunit. Generally, the quantity of catalyst recycled to the quanity ofcatalyst passed to transfer line 8 is about 10:1.

A gaseous efiluent comprising treated naphtha, hydrogen and relativelylow boiling hydrocarbons is removed from separator 43 through conduit 49and is passed through a condenser 50, a conduit 51. to an accumulator52. Any catalyst dust in the efiiuent may be removed prior tocondensation by conventional means, as previously discussed. Condenser50 may comprise a single or series of condensation units similar to unit22 which reduce the temperature of the eiiluent to about 100 F. andcondense the naphtha. Treated naphtha is withdrawn as a product of theprocess from accumulator 52 through conduit 56 and is generallystabilized and rerun before blending into finished gasoline. The qualityand characteristics of the stabilized and rerun naphtha prodnet areshown in Table IV below:

Sulfur, wt. percent 0.015 Nitrogen, wt. percent 0.06

Uncondensed vapors comprising hydrogen and relatively low boilinghydrocarbons are removed from accumulator 52 and a portion subjected toconventional recovery processes to recover gasoline boiling rangeconstituents, and another portion is passed through conduit 53 asrecycle to the process. This gaseous material is recycled throughconduit 53 to conduit 31 and, in some instances, a portion may be passedthrough conduit 57 to conduit 4 and thence to hydrodesulfurizationreactor 9.

The quantity of materials obtained as net product of the process fromhydroformer 4i, excluding recycle, is approximately as follows:

Make gas (Ci-C3 and H2)-9,100,000 standard cubic feet per day 100% C4gasoline8,350 barrels per stream day 400" F. polymer200 barrels perstream day The make gas is rich in hydrogen and it may, therefore, besubjected to fractional separation so as to supply a part of thehydrogen consumed in the hydrodesulfurization step.

Ordinarily, some regeneration of the catalyst is necessary. It ispreferred, therefore, to remove the catalyst for this purpose fromconduit or standpipe 19 through conduit 58, or alternatively oradditionally from conduit 47 by means not shown, and pass it to aconventional regeneration zone 5). In regeneration zone 59, the finelydivided contact material is contacted with oxygen or air at m elevatedtemperature between about 700 F. and about 1200 F. to burn offcarbonaceous deposits thereon. The hot regenerated catalyst may be usedto supply a portion of the heat for the hydroforming reaction. Thecatalyst may also be reduced, if desired, without departing from thescope of this invention. Fresh catalyst may be introduced into thesystem at the same point as the regenerated catalyst.

Although the make-up hydrogen for the process, as described, isintroduced into conduit 4 through conduit 6, this make-up hydrogen maybe introduced conveniently as the stripping gas to stripper 15 throughconduit 18. In such a case, the stripping gas is passed by means notshown from stripper 15 into separator 12 where it is admixed with theeffluent from hydrodesulfurization zone 9. The make-up hydrogen is thenultimately recycled through conduit 37 to feed line 4. The stripping gasmay be heated to aid in stripping and in maintaining conditionsconducive to the removal of deposits by reactions, such as hydrogenationor oxidation.

Specific reaction conditions have been stated for the example describedin connection with the process of Figure 1. These particular operatingconditions should not be construed as unnecessarily limiting to thepresent invention. Both the hydrodesulfurization reaction and thehydroforming reaction may be carried out over a wide range of operatingconditions. The temperature employed for hydrodesulfurization may be anytemperature between about 500 F. and about 1000" F., preferably betweenabout 700 F. and about 925 F. The temperature of reaction ofhydroforming may be between about 850 F. and about 1050 F., preferablybetween about 875 F. and about 950 F. and, in general, higher than thetemperature employed for the hydrodesulfurization step. The pressureemployed for the hydrodesulfurization step is between about 200 andabout 4000 pounds per square inch gage, preferably between about 300 andabout 1200 pounds per square inch gage. The pressure for thehydroforming step is between about 100 and about 1000 pounds per squareinch gage, preferably between about 200 and about 750 pounds per squareinch gage. Since the catalyst is circulated between thehydrodesulfurization zone and the hydroforming zone, it is preferred tooperate both of these reaction zones at substantially the same pressure.Thus, in the type of operation described with respect to Figure 1 inwhich catalyst is circulated between the reaction zones, it is preferredto use substantially the same pressure in each of the reaction zoneswithin the range of about 300 to about 750 pounds per square inch gage.When catalyst is not circulated between the reaction zones the use ofdifferent pressures in the reaction steps is feasible and, under suchcircumstances, relatively higher pressures are employed in thehydrodesulfurization step than in the hydroforming step, but within therange previously stated. It is preferred to use relatively highpressures, particularly in the hydrodesulfurization step, in order tominimize the formation of coke.

In the hydrodesulfurization step, the mol ratio of hydrogen to chargingstock is generally between about 1:1 and about 30:1, preferably betweenabout 2:1 and about 9:1. In the hydroforming step, the mol ratio ofhydrogen to naphtha feed charged from the hydrodesulfurization step isless than that employed in the hydrodesulfurization and is generallybetween about 1:1 and about 8:1.

The space velocity in weight of liquid charged per hour per weight ofcatalyst in the hydrodesulfurization reaction zone is between about0.3:1 and about 15:1, preferably between about 1:1 and about 5:1. Thespace velocity in similar terms for hydroforming is between about 02:1and about 5 :1, preferably between about 0.3:1 and about 2:1. r

Variations in the operating conditions between the reaction zones may beachieved in conventional manner. Desired temperatures may be obtained bythe method to be discussed hereinafter. Variations in mol ratio ofhydrogen to naphtha charge can be achieved by appropriate recycle ratiosand the particular quantity and location of the introduction of make-uphydrogen. Space velocity is controlled by the rate of feed chargedreases or therate"ofcirculation ofcatalyst to the resiie'cave reanemones;

The gas recycle rate to etrlh of the reaction zones' 'is between about2000 and about 15,000 standard cubic feet per barrel of naphtha charged,preferably between about 4000 and about 10,000 standard cubic feet perbarrelof naphtha charged.

Iugeneral, the catalysts used in'the process of Figure 1 comprise theoxides of the metals of the left-hand column of group Vi of the periodictable, particularly chromium, molybdenum, and' tungsten are preferred,but other metallic oxides and other metallic compounds, particularly theoxides of the metals of the left-hand column of groups 1V and V of theperiodic table, such as titanium, serium, thorium and vanadium may beused. The sulfides as well as the oxides of the various metals may beemployed, if desired. The metals and their sulfides of group VIII of theperiodic table may also be used, such as platinum, palladium, nickel andcobalt. While these catalytic oxides and sulfides can be used alone'; itispr'eferable to use them on a suitable supporting material, such asmagnesia or alumina, particularly an activated alumina or an alumina gelsupport. In general, catalytic oxides or other catalytic'co'mpounds arepresent as the minor constituents of the over-all catalyst mass, usuallyfrom 1 to 40 per cent by weight of the total catalyst mass, includingthe support. It is also Within the scope of this invention to useamixture of catalysts; Preferred catalysts comprise cobalt molybdate,molybdenum oxide, tungsten sulfide, tungsten nickel sulfide, tungstenmolybdenum sulfide, unsupported" or supported on activated alumina oralumina gel. When the catalyst is not circulated between thehydrod'esulfurization zone and'the' hydroforming zone, two dilierentcatalysts may be employed in each ormese zones. For example, cobaltmolybdate'may be employed as a catalyst for the hydrodesulfurizationreaction and" molybdenum or chromium oxide may be employed as thehydroforming' catalyst.

Preferably, thefinely divided catalyst of this invention contains nomore thana minor proportion by weightof materialwhose particle size isgreater than about 250 microns; The greater proportion of the catalystmass preferably comprises a material whose particle" diameter is smallerthan 100 microns, including at least 25 per cent of a material having aparticle size smaller than 40 microns. An example of a desirablepowdered catalyst is one which comprises at least 75 per cent by weightof amaterial smaller than 150 microns and atleast'25 per cent by weightof a material smaller than 40 microns;

Althrough- Figure 1 has been'described with reference to an upwardlyflowing gaseous stream of reactants and catalyst through therespectivereacti'on zones, it'sh'ould be understood that the catalystand reactants may flow together downwardly, horizontally, circularly oreven angularly through a reaction zone at a relatively. high velocitywithout'departing from the scope of this inyention. However, it hasbeenfound that by upward flowing of the gases througha substantiallyvertical reaction zone the entrainment of the catalyst and thevresidencetime thereof may be controlled conveniently and accurately and thetendency fonsegregation and stratificationof e ata yst s n m edt Ex n sfe dist k s ch i estq l I sane, may b rqdu din esy 'n e wee t e such asinto conduit 31 by means not shown, Steam may also be'used to replace'atleast a portion of .the added hydrogen in the process. Thus, forexample; steam may be introduced into conduit 4' through conduit 6 andthe steam reacts with the sulfur in hydrodesulfurizer 9 to producehydrogen sulfide. k

According to this invention the exothermic heat liberated by thehydrodesulfurization reaction is transferred to the endothermichydroforming reaction by indirect heat exchange between the reactionzones with the use 10 ofa eosiaiitvaiserceureressae me'reauides s i i'de' substantially: isothe'ir'riial conditio'hs v tier: suitablevaporizable liquidmedium; sueh'aeme cury, is' passednem-reserveir 61'thibug'h duit'62 w heatexchangei" 63'.' The liqtiidgwliich iii theexampleof'Fi'gure'lis niercury, -is vaporized in net chan'ger 63 WEissuitabl'ydisposed inhydrbdesulftinzer 91in heat ex' ang'e'with thereactantsi Thepressureof theimercuiy lsheld' at about 40'poufids per" aua ch gage, of which ressure thevapo'rization teriip sane is about 825F. in he'at exchanger 631 As"preytou'sly'ciis' cussed,the"hydrhdesuliurization reaction oftheexlaiiipl of Figure'l' liberatesabout 161700;00013'1 s'p i hour which iii tuin'evaporates about 141,000pounds" of n'iercury per hour. The evaporation of the mercury abstifbs,the exothermicheat of reaction of reactor 9 by latent heat ofvaporization,

The mercury vapor thus evaporatedis removed from heat exchanger 63through conduit 64' and returnedwe reservoir 61. Reservoir 61*ismaintained atthe op eiat ing pressure of about 40 pounds persquare'in'ch gag-e1. Mercury vapors are re'movedfromfl reservoir 61through conduit'67 and passed through conduit 71 to a two sta'geturbo-compressor 72 in which the vapor is compressed to a pressure ofl20pounds' per sc' uare inch gageand the temperature is'raised toabout1175 F. Two stagesof compression with interstage-cooling' are usedbecause the low specific heat of'the' mercury vapor results iii a" tendency towards excessive temperature rise during compression:The'coinpre'ssed' mercury at'avte'rnperature ofabout 1-175' .F. andabout pounds per square inchgage is passed fromicom'p'ressor '72'throughconduit 74 to heat exchanger, 76; Mercury vapor flows through heat exchanger 76in indirect exchange with the reactant gases'in hydroformer 41and supplies the 16,200,000131; u.s"per hour of'heat required'forhy'dr'ofor'mer 41'. The re"- moval of heat from the mercury vapor at the120 pounds pressure 'results'in its condensation: at approximately 950F. in heat exchanger 76; Liquid mercury is removed from heat' exchanger76 through conduit 77 and is expandedthrough expansion valve 78 intoconduit 64. Duringexpansion of the mercury in'valve 78; a small portionof the mercury is again vaporized. The mercury inconduit 64 is recycledto reservoir 61.

Since the hydroforrnihg-heat duty is' slightly less'than that removedfrom the hydrodes'ulfurization reactor and alsobe cause of the' heatadded to the compressed vapor as work, the quantity of vapor required tobe condensed in exchanger 76' isless' than that vaporized in heatexchanger" 63, this amount beingabout 133,000 pounds per hour; Theremainder or about'8,000'po'unds" per hour of mercury vapor plus' theamount resulting from the dashing of'the l20pounkl'sper square inch gagemercury condensate through valve' 78 are passed from reservoir 61through'c'onduit 67 to condenser 68 to remove'e'xcess heat and' theresultingcon'densate returned to' reservoir 61 ,-.as" shown; in order tobalance thesystem.

.Theiopera'ti'on of compressor 72, which has an overall efiiciency'ofabout'70per' cent; consumes" about 800 horsepower; Since the heat"equivalent of one horsepower hour is about 2,545 B; t. u:s, 16,200,000B; t.'u'.s per hour are recovered and made available at the highertemperature level required for hydroformer 41 at an energycostcorresponding to only about 2,040,000 B: t. u;s per hour.

As shown iri'elevationl'the' condensate from condenser 68 is p'assedtoreservoir 61 by the static head of the condensate in the'conduitconnecting condenser 68" and reservoir 61; Should this static head beinsufficient to pass the condensate from condenser 68 to reservoir '61,the required amount of vapors to balance the system-may be removed fromthe downstream side of compressor 72 through conduit 73' and passed tocondenser 68 at an elevated pressure.

While mercury is' used as" a thermodynamic fluid or heat transfer mediumin the example of Figure 1, any stable high-boiling liquid may be usedwithout departing from the scope of this invention. Preferably, theboiling point of the heat transfer medium should be about 800 F. orlower in order to avoid sub-atmospheric pressures on the heat inputside. Another suitable high-boiling heat exchange fluid for use in thepresent system is a close-cut fraction of the hydroformer polymer itselfproduced by the chemical process shown in Figure 1. This high-boilingpolymer is highly aromatic and, therefore, heat stable. Although heatexchange with reactor 9 is shown as downfiow and with reactor 41 asupflow, heat exchange with the reactors may be both upflow or bothdownfiow without departing from the scope of this invention.

The indirect transfer of heat from the low temperature exothermicreaction to the high temperature endothermic reaction in the mannerdescribed is highly advantageous since it minimizes the amount ofpreheating necessary for the hydroformer feed stock. As would beobvious, the use of high preheat temperatures characteristic of theprior art results in undesirable thermal decomposition and also makes itimpossible to obtain isothermal conditions in the reaction zone.

The high velocity system shown for both hydrodesulfurization andhydroforming has certain particular advantages which make such a systempreferable to conventional fluidized dense phase catalyst operations.Hydrodesulfurization and hydroforming are not like many reactions, suchas the hydrogenation of carbon monoxide to produce organic compounds,which reactions may be effected at low conversions per pass withrecycling of the reactants to increase the over-all conversion. In thetype of reactions discussed in this application, high conversions mustbe achieved, if at all, in a single pass without recycling. In fluidizeddense phase operations due to the low gas velocities, high conversionsper pass are difficult. This difficulty arises from the internal circulation of reactants and catalyst in the reactor resulting in ampleresidence time for some material but insufiicient residence time forother material. The high velocity system or concurrent flow of catalystand reactants obviates the above difficulties. In the high velocitysystem both the residence time of the catalyst and reactants are undercontrol inasmuch as there is a minimum of internal circulation.Substantially complete removal of sulfur can be obtained in a singlepass by the high velocity technique.

Numerous other advantages are apparent with the high velocity systemwhich cannot be obtained readily with a conventional fluidized densephase operation. In the high velocity system the residence time of thecatalyst in the reaction zone is relatively short as compared to thedense phase type of operation and between each pass of the catalystthrough the cycle of the high velocity system the catalyst is strippedof tarry and carbonaceous deposits and occluded gases. This frequentstripping of the materials maintains the catalyst at its maximumactivity and increases the life of the catalyst; in some caseseliminating entirely regeneration of the catalyst. The high dispersionand rapid movement of the catalyst in the high velocity type ofoperation also minimizes the sticking together of the catalyst particlesas a result of the accumulation of tarry and carbonaceous depositsthereon. The dispersed condition also accounts for a more smoothreaction and minimizes the chance for overheating frequentlyaccompanying exothermic reactions. The high velocity system also ischaracterized by A full control of the concentration of the catalyst inthe reaction zone, this concentration being a direct function of thevelocity of the reactants and the loading rate of the catalyst into thegas stream. This is not true of the dense phase operation Where theconcentration is essentially a function of the velocity only. Still afurther advantage of the high velocity system is that at the highervelocities characteristic of thissystem heat transfer between the wallsof the heat exchange means and the reactants is very efficient as thegas film resistance is minimized.

With regard to the heat transfer mechanism between the two differentreaction zones, this method of heat transfer may be applied tostationary bed reactors and fluidized dense phase reactors as well asthe high velocity type of reactor. However, as previously stated thehigh velocity system is particularly adaptable to this type of indirectheat transfer since it has been found that the rate of heat transfer isexceptionally high.

In addition to the hydrodesulfurization of a light naphtha or gasolinefraction, a gas oil or kerosene frac: tion may also be simultaneouslyhydrodesulfurized. While the hydrodesulfurization of a naphtha fractionhas particular utility in the present invention, since this fraction issubsequently subjected to hydroforming and the heat of exothermichydrodesulfurization may be transferred to the hydroforming reaction, itis within the scope of this invention to simultaneously desulfurize gasoil or kerosene and to transfer the heat from both the desulfurizationof the naphtha fraction and the gas oil fraction to the hydroformingreaction of the naphtha fraction. For such a set-up two or more heatexchangers in parallel may be employed for the hydrodesulfurizationreactions using separate reactors. These heat exchangers are connectedin substantially the same manner as described with respect to Figure 1and the heat liberated by the desulfurization reactions is absorbed by astream of the same heat transfer medium and passed to a reservoir wherethe vapors are compressed and then passed to the heat exchanger of thehydroforming reaction zone.

Although the invention has been described with respect tohydrodesulfurization of a suitable hydrocarbon fraction, selectiveoxidation of the sulfur with oxygen or air may be carried out insubstantially the same manner as described with respect to thehydrodesulfurization reaction. The primary differences are thesubstitution of oxygen for hydrogen in the hydrodesulfurization reactorand the removal of sulfur dioxide from the efiluent of thedesulfurization reactor by distillation rather than by extraction orabsorption. The same catalyst may be employed for oxidation of thesulfur as with the hydrodesulfurization reaction described hereinbefore.However, in oxidizing the sulfur it is preferred to use vanadiumpentoxide on silica gel for the sulfur removal stage.

In the previous discussion of the invention the contact material wasreferred to as a catalytic material, but it is within the scope of thisinvention thatthe contact material may comprise a substantial proportionof relatively inert finely divided material, such as acid treatedbentonite, finely divided charcoal, powdered silica gel, etc. The use ofinert contact material is addition to the catalytic material aids infiuidization of the catalytic material which may, under certaincircumstances, be difficultly fluidizable. In using an inert material incombination with the catalyst a proportion of the inert material tocatalytic material is generally between about 1110 to about 1:1 byweight.

The upgrading or improvement in quality of petroleum distillates by thecombination of hydrodesulfurization and hydroforming in separate steps,as shown by Figure 1, has numerous advantages. Most large refineries,even those adequately equipped with catalytic capacity, pro duce asubstantial amount of thermally cracked gasoline, generally as much at25 per cent or more of the total gasoline output. The sources of thesestocks are generally visbreaking or coking, cycle stock cracking andstraight run naphtha reforming operations. By todays standards thequality of such fractions is comparatively low, averaging perhaps C. F.R. M. octane, with poor stability and tetraethyl lead response. Inaddition, these thermally cracked fractions are frequently high insulfur greases 13 compounds which are not removable by conventionalmethods. In the past such fractions have been drsposed ofor utilized bydilution, i. e., blending with higher quality gasolines from catallyticcracking or hydroforming. However, in view of present day standards ofincreased quality and the amount of increase in sour crudes, theutilization of such thermally cracked fractions is becoming a criticalproblem to many refiners.

Hyd'roforming alone is not a satisfactory solution to the above problembecause of the low yield of product of suitable octane number, high cokeyields, the requirementfor corrosion resisting materials ofconstruction, and thepollution of the surrounding atmosphere with sulfurcompound. The combination process of hydrodesulfuriza'tion andhydroforming has, therefore, become of great importance and the need forimproving the efliciency and operability of such processes is'appareht.The combination process as described with respect toFigure 1 over; comesessentially all of the above difiicultie's'and 'is'co'mmerciallyfeasible, particularly when the use of a" high velocity concurrentsystem is employed.

Aspre'viously stated, Figure 2 of the'drawin'gisa gran'nnaticillustration of multi-stage heat exchange which maybe convenientlyemployedon' hydrodesul fu'rization reactor 9 and hydroforming reactor41"of Figure 1. The use of the multi-stage heat exchange on thefreactors, "as shown, enabl'es'more close temperature control of "theprocess. For example, the heat release or heat absorp tion'at variouspoints in the reactors may vary: and, as a result, the quantity of heatto be removed or added must be correspondingly varied, whichtemperatureicon' tr'ol by a single heat exchanger is difiicult;Purtherrit may be desirable to maintain different temperature levels atclifierent points in the reactors which can' be achieved best by the useof multi-stage heat'exchan'ge. The air rangernent of apparatus in Figure2 is a suitabl'e'set-up'for maintainingditlerent temperature levels T1,I2,"[T3' an'd T4 in reactors 9 and 41 of Figure 1. It is" to be'u rrder}stood, however, that the same set-up may be used when the quality ofheat to be removed is di'fierent'at different points in the reactors,even when T1 and' Tz, and and T4 are to bemaintained at substantiallytheisame levels, respectively. According to Figure: 2, temperature T1 ismaintained substantialy constant by maint :a pressure P1 in reservoir101 such that the heat exchange fluid boils at a temperature suitablyrelated 'to' 'li :Ifhe liquid heat "exchange medium, such as merc'ury, vassed from reservoir 101 under pressure P1 through conduit :2 to heatexchanger 103 where the liquid boils at 'a tempera} ture correspondingtopressure P1. The vapors ofthe heat exchange medium are removed from heatexchanger 103 through conduit 104 and returned to, reservoir 101*.Vapors are"rein'oved from reservoir 101 through'conduit 106 to be passedto the heat exchange unit ofhy'drofornff ingreactor 41 as to bediscussed further here'ipaftet: In another portion of reactor 9,temperature'Tz is main} tamed by adjusting the pressure P2 ofreservoir'107'at the" required value. Liquid heat exchange medium iswithdrawn" from reservoir 107 through conduit" 108" and passed to" heatexchanger 109. In heat exchanger 109, the liquid heat exchange medium isvaporized at a 'tm p'erature corresponding to pressure P2. Vapdrotish'atexchange medium is withdrawn from heat exchanger 109 through conduit 111and passed to reservoir 107. Va porous heat exchange medium is withdrawnfrom reset" voir" 107 through conduit 112 to be passed to theheatexchange unit of hydroforming reactor 41 as to be dis cussedhereinafter.

,v The exothermic heat of hydrosulfurization unit 9 is made available tohydroforming unit 41 by passing the vaporized heat exchange medium fromconduit 106 to compressor 116. In compressor 116, the vapors arecompressed in a similar manner as described with respect to Figure l topressure P; which corresponds to the pressure necessary to maintain atemperature of the 14 Heait' exchange martinis suitably related: toteniperature Ts at"thatparticularpoint"showh' in ireactor 41.1 ThecOmprjess edyapors' are passed from compressor 116 through conduit 117to heafexchanger 1118* ln'heatexchanger 118 the compressed vapors'arecondensed at a temperature corresponding to pre's'su're' P giving'f uptheir latent heatofconHensafionto the reaction mixture iri'theupperportion of'react'or 411 Condensedheat exchange medium is passed'fromheat exchanger 118 through conduit 119 to conduit 121 andthen returnedto reservoir 101 after expansion through an expan's'iofivalve in'line119.1 Similarly,- v'aporous' hea't exchange medium from reservoir 107 ispassed:throughconduit to compresset 123. The vaporous heatexchah'gemediunitiscornpressed to a pressure P4 in"co' inpressor 123. Theco'mpre'ss'ed'vap'ors are passedf rbm compressor 123 through conduit 124 to'heatexchanger 126 i" Ih he'at exchanger 126, the compressed vaporsarecorrd'ensed at a' temperatu're corresponding to pressure" P4 and giveup their latent heat of condensationdo thereactionunixture 'in'theupperportion of reactor 411 Condensed heatexchange medium is passed from heatexchanger 118 through conduit" 1'19 to'conduit 121' and thenreturned'to' reservoir 101*after expansion throughfanlexpansion valveinjline 119; Similarly, vaporous heat exchange'medium'from'reservoir 107is'passe'd through c'on'duit to compressor 123. The vaporous heatexchange mediumis compressed to a ssure P r'in compressor 123;Thecornpressed vapors are passed' from compressor 123 throughconduit 124to heat exchanger 126'. In" heat exchanger; 126, i the compressed vaporsare condensed at"aitemperaturecorresponding to pressure P; and give"their latent heat of condensation to the"re'actants in the"lo wer"portion 'ot" hydroformer 4'1.

. Condensate"isp'assed from heat exchanger 126'through conduit'1'27 andan expansion valvetoj reservoir 107.

If the exothermic heat of the 'hydrodes'ulfu'rization reaction effected"in reactor9is" greater thantha't necessary to supply heat to theendothermic reaction eifected' in reactor 41 and to maintaintheternperatu're'a't the desired level or levels thereim'a" portion ofthe" vapors in conduits '106'andj/o'r' 112" is re'r'r'ioved' therefromthrough conduit 131 and passed throughc'onduit 132w condenser 133; Thisportion of the" vapors is" condensed in corn denser 133."Fromcon'dlenser and'cooler' 133" condensate is returned to thesystertithrou'gh' conduits134' and'136. This condensate may be passeddirectly to reservoir 101 or' directly to reservoir 107 or may be splitand'passed in the appropriate proportions to both reservoii's'lttl and107 as shown. In some'instances the amount of heat available from.hydrodesulfuriz'ation' reactor 9" or from any' exothermic reaction whichmay be eliected therein and the compression of the "heat exchange vapormay be insuflicient to'supplytherequitedainount of heat to the reactioneffected in reactor 41. In such case the additi'onal'heat requiredobtained by passing condensate frorn'conduits 119 and/or127'throughconduit 138 and conduit 139 to vaporizer 141. In vaporizer" 141', therequired amount of'm'ake-up heat is added to the system to effect thevaporization of that portion of the condensate removed throughcondu'it138. The vapors produced in heater 141'are passed through conduit 136and returned to the system through conduit 121 and/or conduit 127.Thse'va'p'ors' may be passed directly to reservoir 101, or reservoir107, or may be split and passed in appropriate proportions to reservoirs101 and 107, respectively.

The above arrangement of" apparatus is an eiTectivc manner formaintaining different temperature levels in the exothermic andendothermic reactors as Well as maintaining a constanttemperaturethroughout the reactors where the'heat liberated or absorbed isdifferent for difierentlocatio'ns within the reactors. Numerouscombinations are available in the arrangement shownin Figure 2;Temperature T1 may be higherthantemperature T2 and temperature'ls may behigher than temperature T4 in" one combination. In; another combination,T1 and T a may be at the same temperature level and T3 and T4 may be atthe same but higher temperature level, but the heat duty at T2, T3 andT4 may be different as the result of different rates of reaction at eachlocality. In still another combination, T2 may be greater than T1 and T4may be greater than T3. It is also within the scope of this invention tohave the temperature of T1 less than the temperature of T2 but at thesame time having temperature T3 greater than temperature T4. In each ofthe above combinations pressures P1, P2, P3; and P4 must be adjustedaccordingly.

The arrangement of apparatus and heat exchangers of Figure 2 areparticularly adaptable Where a second exothermic reaction (not shown) isbeing effected in the over-all process, from which additional heat maybe available for heating up the reaction in the endothermic reactor,such as in reactor 41. For example, a heat exchanger may be inserted inthe catalyst regeneration zone, such as regenerator 59 of Figure 1, andconnected to the heat exchange system in a manner similar to heatexchanger 1193 or 109. The regeneration of the catalyst is largelyachieved by oxidation of carbonaceous deposits, which reaction isexothermic. The arrangement shown in Figure 2 may also be employed whena gas oil or reduced crude is being hydrodesulfurized in a separatereactor from which heat may be made available by the system disclosed inFigure 2.

The multi-stage heat exchange system of Figure 2 is particularlyadaptable where multi-stage hydrodesulfurization is effective as theheat liberated in each stage of the hydrodesulfurization may bedifferent and, therefore, to take advantage of this liberated heat it isnecessary to employ separate heat exchangers for each zone and pass theheat thus recovered to the endothermic zone, such as endothermic zone 41of Figure 1. Figure 3 is a diagrammatic illustration of an arrangementof apparatus for effecting multi-stage hydrodesulfurization. Multi-stagehydrodesulfurization is particularly advantageous, as previously pointedout, since in fluidized processes it is substantially impossible toobtain maximum or high degree of desulfurization unless some provisionis made for minimizing internal circulation of catalysts and reactants.Fluidized hydrodesulfurization can best be effected by the concurrenthigh velocity process, and maximum efliciency even of this process isachieved by effecting the desulfurization reaction in stages. Accordingto this modification of the invention, a suitable naphtha feed stockcontaining sulfur is passed through conduit 151 to heater and vaporizer153. Make-up hydrogen is introduced into conduit 151 through conduit152. vaporized naphtha is passed through conduit 154 to a first stagehydrodesulfurization reactor 156. Contact material is introduced intothe stream of vaporized naphtha and hydrogen in conduit 154 throughconduit 178. The reactants and contact material pass upwardly throughreactor 156 in contact with heat exchanger 157. The efiluent from thefirst stage hydrodesulfurization containing entrained catalyst is passedthrough conduit 159 to catalyst separator 161. In catalyst separator16]. catalyst is separated from naphtha, hydrogen sulfide and hydrogenand settles into stripper 162. Separator 161 and stripper 162 aresubstantially the same as the corresponding separator and stripperdescribed in Figure 1 of the drawings. A suitable stripping gas, such ashydrogen or steam, is introduced into stripper 162 through conduit 163.Stripped catalyst is removed from stripper 162 through a conduit orstandpipe 164.

The partially clesulfurized naphtha and hydrogen are removed fromseparator 161 through conduit 168 and passed to the second stagehydrodesulfurization unit 169, in which desulfurization of the feedstock is completed. Stripped catalyst from the first stage reaction isintroduced into conduit 168 through conduit 166, as shown. A portion ofthis stripped catalyst may be recycled to the first stagedesulfurization unit 156 through conduit 167,

if desired. Vaporized naphtha, hydrogen and entrained catalyst arepassed upwardly through second stage reactor 169 in contact with heatexchanger 171. The effluent from the second stage hydrodesulfurizationunit 169 is withdrawn through conduit 172 and passed to catalystseparator 173 in which entrained catalyst is separated from the reactioneffluent. The entrained catalyst settles second stage desulfurizationunit may be recycled to conduit 168 through conduit 179, if desired.

The naphtha effiuent substantially free from organically combined sulfuris removed from separator 173 through conduit 181 and passed throughconduit 182 to a conventional condensation stage 183. The condensationstage 183 may comprise a series of condensation units at the samepressure or at successively lower pressures. Condensate is passed fromcondenser 183 to accumulator 184. From accumulator 184, the desulfurizednaphtha is withdrawn through conduit 186 as a product of the process.The fraction recovered from conduit 186 may be subjected to furthertreatment, such as hydroforming. Uncondensed vapors comprising hydrogen,hydrogen sulfide, methane, ethane and other relatively low boilinghydrocarbons are removed from accumulator 184 through conduit 187 and aportion of this vaporous stream may be vented to the atmosphere toprevent the build-up of ethane and methane in the process. A majorproportion of the vapor stream is recycled after removal of hydrogensulfide as in Figure 1, through conduit 188 to feed line 154 to supplythe added hydrogen to the process. A portion of the vapor streamcontaining hydrogen may be passed through conduit 189 to conduit 168 inorder to control the hydrogen content of reactor 169 independent ofreactor 156, if desired.

It is preferred to introduce fresh catalyst into the system throughconduits 192 or 193 between the first and second stage of the process. Aportion of the effluent withdrawn from the first stage of the processmay be by-passed and passed to condenser 183 through conduit 191, ifdesired.

In the arrangement of apparatus shown, heat exchanger 157 mayconveniently constitute heat exchanger 109 of Figure 2 and heatexchanger 171 may constitute heat exchanger 103 of Figure 2. Heatexchanger 157 and 171 may also be connected in series as a singleexchanger and connected in the system like shown in Figure 1.

In general, the reaction conditions and method of operation are similarto those described with respect to the hydrodesulfurization unit 9 ofFigure 1. It is possible and often desirable to maintain the first andsecond stage desulfurizationat different conditions when using amultistage hydrodesulfurization process. Thus, it is desirable tomaintainthe second stage hydrodesulfurization reaction at a temperatureat least 50 F. higher than the first stage and preferably above 800 F.,but within the range disclosed with respect to reactor 9 of Figure 1. Atsuch higher temperature in reactor 169, desorption, vaporization anddestructive hydrogenation of carbonaceous deposits and tars depositedupon the catalyst is effected and a material increase in the length ofthe catalyst life is observed. Under optimum operating conditionsregeneration of the catalyst may be eliminated entirely. It may also bedesirable to maintain the hydrogen content in the second stagehydrodesulfurization higher than the first stage. In the preferredmodification of Figure 3, the hydrogen content in terms of tool ratio ofthe first stage is maintained approximately half that of the secondstage but within the range as disclosed with respect to reactor 9 ofFigure 1. It may be also desirable to maintain'the reaction pressure ofthe second stage hydrodesulfurization at a higher pressure than in thefirst stage, In such a modification, the hydrogen content of thereaction mixture may be substantially the same in both stages, but thepartial pressure of the hydrogen may be substantially greater in thesecond stage. Increase in pressure has a similar effect as an increasein temperature and aids in the destructive hydrogenation of carbonaceousand tarry deposits from the catalyst. Under the preferred operationconditions for the two-stage process of Figure 3 between about 50 andabout 70 per cent of the organic sulfur is converted in the first stageand the remainder of the sulfur up to about 90 to 100 per cent isconverted in the second stage. Although only two stages of operation areshown for the modification of the invention of Figure 3, three and fourstages or more may be employed without departing from the scope of thismodification.

A similar operating technique of multi-stage operation may be employedto effect hydroforming. In hydroforming by multi-stage process theoperation is carried out substantially the same as that described withrespect to Figure 3. It is particularly desirable in the case ofhydroforming to carry out the second or last stage of the process at ahigher temperature than the first stage in order to completearomatization. A temperature increase of at least 25 F. is preferred inthe case of hydroforming in a multi-stage process.

It is within the scope of the present invention to employ a condensationstep between the first and second stage of the multi-stage processdescribed with respect to Figure 3. Condensation between stages may bedesirable in order to have independent control over the quantity andquality of recycle gases to the respective stages. It may also bedesirable to remove the hydrogen sulfide between stages and,consequently, condensation of the effluent from the first stage will benecessary as described with respect to Figure 1.

In another modification of the present invention, thehydrodesulfurization step may employ the high velocity technique asshown in Figures 1 and 3 followed by a hydroforming step conducted insuch a manner as to maintain the catalyst in the so-called pseudo-liquiddense phase condition. This particular type of operation may bedesirable in order to obtain increased contact times in the secondhydroforming reaction step, whereas the high velocity type of operationmay be preferable for the hydrodesulfurization step because of therelatively short contact times permissible.

The drawings of the present invention diagrammatically illustratevarious modifications of the invention. Various pieces of equipment,such as compressors, storage vessels, condensers and coolers,fractionation and extraction equipment, etc., have been eliminated fromthe drawings as a matter of convenience and clarity and their use andlocation will become apparent to those skilled in the art.

Although the invention has been described with particular reference tohydrodesulfurization and hydroforming, the invention may be applied toother combination processes employing both exothermic and endothermicreactions in which the exothermic reaction is carried out at a lowertemperature than the endothermic reaction. For example, the inventionapplies to synthesis gas making and synthesis, hydrogenation andhydroforming, synthesis reaction and catlaytic cracking or reforming ofa product thereof, synthetic phenol(l) chlorination of benzene and (2)hydrolysis of chlorobenzene, and many other chemical processes.

Having described my invention, I claim:

1. Apparatus for the indirect heat exchange with a heat exchange fluidbetween a plurality of interconnected successive reaction chambers whichcomprises in combination a first elongated reaction chamber, a secondelongated reaction chamber, two separate heat exchange means havinginlets and outlets therein adapted for heat exchange with said firstreaction chamber, two other separate heat exchange means having inletsand outlets therein adapted for heat exchange with said second reactionchamber, the aforesaid heat exchange means being positioned at spacedintervals longitudinally along the elongated reaction chambers, twoaccumulators for heat exchange fluid, conduits connecting the inlets andoutlets of corresponding heat exchange means of said first and secondreaction chambers with each accumulator toform two separate heatexchange circuits between said reaction chambers, each of said heatexchange circuits comprising, in addition to said accumulator and saidheat exchange means, a conduit from the lower portion of saidaccumulator to the inlet of the heat exchange means in the firstreaction zone, a conduit from the outlet of the heat exchange means inthe first reaction zone to said accumulater, a conduit from the upperportion of said accumulator to the inlet of the heat exchange means inthe second reaction zone, a compressor in said last-named conduit, aconduit from the outlet of the heat exchange means in the secondreaction zone to said accumulator and a pressure reducing means in saidlast-named conduit, means for passing heat exchange fluid from each ofthe above heat exchange circuits to a separate and fifth heat exchangemeans, and means for returning fluid from said separate and fifth heatexchange means to said heat exchange circuits.

2. Apparatus for the indirect heat exchange with a heat exchange fluidbetween a plurality of interconnected successive reaction chambers whichcomprises in combination a first elongated reaction chamber, a secondelongated reaction chamber, two separate heat exchange means havinginlets and outlets therein adapted for heatexchan'ge with said firstreaction chamber, two other separate heat exchange means having inletsand outlets therein adapted for heat exchange with said second reactionchamber, the aforesaid heat exchange means being positionedlongitudinally along the elongated reaction chambers, two accumulatorsfor heat exchange fluid, conduits connecting the inlets and outlets ofsaid heat exchange means of said first and second reaction chambers witheach accumulator to form two separate heat exchange circuits betweensaid reaction chambers, each of said heat exchange circuits comprising,in addition to said accumulator and said heat exchange means, a conduitfrom the lower portion of said accumulator to the inlet of the heatexchange means in the first reaction zone, a conduit from the outlet ofthe heat exchange means in the first reaction zone to said accumulator,a conduit from the upper portion of said accumulator to the inlet of theheat exchange means in the second reaction zone, a compressor in saidlast-named conduit, a conduit from the outlet of the heat exchange meansin the second reaction zone to said accumulator and a pressure reducingmeans in said last-named conduit, means for passing heat exchange fluidfrom one of the above heat exchange circuits to a separate and fifthheat exchange means, and means for passing fluid from said separate andfifth heat exchange means to the other of said heat exchange circuits.

3. Apparatus for transferring heat between separate interconnectedreaction chambers which comprises in combination two separateinterconnected reaction chambers, two separate heat exchange meanshaving inlets and outlets therein in indirect heat exchange relationshipwith each of said reaction chambers, two accumulators for heat exchangefluid, conduits connecting the inlets and outlets of each of said heatexchange means in pairs with each of said accumulators to form twoseparate heat exchange circuits between said reaction chambers, each ofsaid heat exchange circuits comprising, in addition to said Iaccumulator and said heat exchange means, a conduit from the lowerportion of said accumulator to the inlet of the heat exchange means inthe first reaction zone, a conduit from the outlet of the heat exchangemeans in the first reaction zone to said accumulator, a conduit from theupper portion of said accumulator to the inlet of the 19 heat exchangemeans in the second reaction zone, a compressor in said last-namedconduit, a conduit from the outlet of the heat exchange means in thesecond reaction zone to said accumulator and a pressure reducing meansin said last-named conduit, and means for passing fluid between saidcircuits.

4. Apparatus for transferring heat between separate interconnectedreaction chambers which comprises in combination two separateinterconnected reaction chambers, two separate heat exchange meanshaving inlets and outlets therein for a heat exchange fluid in indirectheat exchange relationship with each of said reaction chamhers, twoaccumulators for heat exchange fluid, conduits connecting the inlets andoutlets of each of said heat exchange means in pairs with each of saidaccumulators to form two separate heat exchange circuits'between saidreaction chambers, each of said heat exchange circuits comprising, inaddition to said accumulator and said heat exchange means, a conduitfrom the lower portion of said accumulator to the inlet of the heatexchange means in the first reaction zone, a conduit from the outlet ofthe heat exchange means in the first reaction zone to said accumulator,a conduit from the upper portion of said accumulator to the inlet of theheat exchange means in the second reaction zone, a compressor in saidlast-named conduit, a conduit from the outlet of the heat exchange meansin the second reaction zone to said accumulator and a pressure reducingmeans in said last-named conduit, a condenser, means for passing vaporfrom said accumulators to said condenser and means for returningcondensate from said condenser to said accumulators.

5. Apparatus for transferring heat between separate interconnectedsuccessive reaction chambers which comprises in combination two separateinterconnected reaction chambers, two separate heat exchange meanshaving inlets and outlets therein for a heat exchange fluid in indirectheat exchange relationship with each of said reaction chambers, twoaccumulators for heat exchange fluid, conduits connecting the inlets andoutlets of each of said heat exchange means in pairs with each of saidaccumulators to form two separate heat exchange circuits between saidreaction chambers, each of said heat exchange circuits comprising, inaddition to said accumulator and said heat exchange means, a conduitfrom the lower portion of said accumulator to the inlet of the heatexchange means in the first reaction zone, a conduit from the outlet ofthe heat exchange means in the first reaction zone to said accumulator,a conduit from the upper portion of said accumulator to the inlet of theheat exchange means in the second reaction zone, a compressor in saidlast-named conduit, a conduit from the outlet of the heat exchange meansin'the second reaction zone to said ac cumulator and a pressure reducingmeans in said lastnamed conduit, a vaporizer, means for passing liquidheat exchange fluid from the outlet of at least one of said heatexchange means of said second reaction chamber, and means for returningvapor from said vaporizer to at least one of said accumulators.

References Cited in the file of this patent UNITED STATES PATENTS

1. APPARATUS FOR THE INDIRECT HEAT EXCHANGE WITH A HEAT EXCHANGE FLUIDBETWEEN A PLURALITY OF INTERCONNECTED SUCCESSIVE REATION CHAMBERS WHICHCOMPRISES IN COMBINATION A FIRST ELONGATED REACTION CHAMBER, A SECONDELONGATED REACTION CHAMBER, TWO SEPARATE HEAT EXCHANGE MEANS HAVINGINLETS AND OUTLETS THEREIN ADAPTED FOR HEAT EXCHANGE WITH SAID FIRSTREACTION CHAMBER, TWO OTHER SEPARATE HEAT EXCHANGE MEANS HAVING INLETSAND OUTLETS THEREIN ADAPTED FOR HEAT EXCHANGE WITH SAID SECOND REACTIONCHAMBER, THE AFORESAID HEAT EXCHANGE MEANS BEING POSITIONED AT SPACEDINTERVALS LONGITUDINALLY ALONG THE ELONGATED REACTION CHAMBERS, TWOACCUMULATORS FOR HEAT EXCHANGE FLUID, CONDUITS CONNECTING THE INLETS ANDOUTETS OF CORRESPONDING HEAT EXCHANGE MEANS OF SAID FIRST AND SECONDREACTION CHAMBERS WITH EACH ACCUMULATOR TO FORM TWO SEPARATE HEATEXCHANGE CIRCUITS BETWEEN SAID REACTION CHAMBERS, EACH OF SAID HEATEXCHANGE CIRCUITS COMPRISING, IN ADDITION TO SAID ACCUMULATOR AND SAIDHEAT EXCHANGE MEANS, A CONDUIT FROM THE LOWER PORTION OF SAID ACCUMLATORTO THE INLET OF THE HEAT EXCHANGE MEANS IN THE FIRST REACTION ZONE, ACONDUIT FROM THE OUTLET OF THE HEAT EXCHANGE MEANS IN THE FIRST REACTIONZONE TO SAID ACCUMULATOR, A CONDUIT FROM THE UPPER PORTION OF SAIDACCUMULATOR TO THE INLET OF THE HEAT EXCHANGE MEANS IN THE SECONDREACTION ZONE, A COMPRESSOR IN SAID LAST-NAMED CONDUIT, A CONDUIT FROMTHE OUTLET OF THE HEAT EXCHANGE MEANS IN THE SECOND REACTION ZONE TOSAID ACCUMULATOR AND A PRESSURE REDUCING MEANS IN SAID LAST-NAMEDCONDUIT, MEANS FOR PASSING HEAT EXCHANGE FLUID FROM EACH OF THE ABOVEHEAT EXCHANGE CIRCUITS TO A SEPARATE AND FIFTH HEAT EXCHANGE MEANS, ANDMEANS FOR RETURNING FLUID FROM SAID SEPARATE AND FIFTH HEAT EXCHANGEMEANS TO SAID HEAT EXCHANGE CIRCUITS.